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A single step non-catalytic esterification of palm fatty acid distillate (PFAD) for biodiesel production
Fuel 93 (2012) 373–380
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A single step non-catalytic esterification of palm fatty acid distillate (PFAD) for biodiesel production Hyun Jun Cho a, Soo Hyun Kim b, Seok Won Hong a, Yeong-Koo Yeo a,⇑ a b
Department of Chemical Engineering, Hanyang University, Haengdang-dong, Sungdong-gu, Seoul, Republic of Korea Chemicals R&D Center, SK Chemicals, 686 Sampyeong-dong, Bundang-gu, Seongnam-si, Gyeonggi-do, Republic of Korea
a r t i c l e
i n f o
Article history: Received 11 October 2010 Received in revised form 26 August 2011 Accepted 26 August 2011 Available online 13 September 2011 Keywords: Biodiesel Non-catalytic esterification Palm fatty acid distillate Water-free reaction
a b s t r a c t In this work, the single step method for non-catalytic esterification of palm fatty acid distillate (PFAD), which is readily applicable to actual production of biodiesel, was investigated. In this method, the esterification reaction is accomplished in a single step by ensuring water-free reaction conditions and the acid value is reduced to below 0.5 (mg KOH/g) which has not been possible in previous methods. The reaction was completed (<0.5 mg KOH/g) within 180 min at relatively high temperature (>250 °C) enough to be above boiling point of water and at moderate pressure (0.85–1.20 MPa) without any catalyst. The effects of temperature, methanol feed rate and pressure on a semi-batch reaction were investigated and the optimal values of these variables were found (temperature: 290 °C, pressure: 0.85 MPa, feed rate: 2.4 g/min). The acid value was reduced from 191.4 to 0.36 (mg KOH/g) just in 180 min at these conditions. From the kinetic study on non-catalytic esterification of PFAD, it was found that the activation energy is 17.74 kJ/ mol and the frequency factor is 2.12 min1. Ó 2011 Elsevier Ltd. All rights reserved.
1. Introduction As an alternative diesel fuel, biodiesel is made out of renewable biomass such as vegetable oils, animal fats, and waste grease. It is well known that the utilization of biodiesel is carbon–neutral and contributable to the reduction of CO2 emission [1,2]. Due to its biodegradability, nontoxic property and low emission profiles, biodiesel has attracted lots of attention and has been considered environmentally beneficial [3,4]. Because biodiesel contains no sulfur but about 11% of oxygen (w/w), we can expect SOx free emission and a significant reduction of pollutant emissions such as unburned hydrocarbons, carbon monoxide and particulate matters compared with petroleum-derived diesel fuel [5,6]. In addition to the environmental benefits, primary advantage of biodiesel is that none- or few modifications of current diesel engine system are needed to adopt biodiesel blends or even pure biodiesel [7]. In general, biodiesel, fatty acid methyl ester (FAME), is efficiently produced from the transesterification of vegetable oil or animal fats or from esterification of fatty acids with short chain alcohols in the presence of homogeneous or heterogeneous alkali- and/or acid–based catalysts [8]. Among catalytic processes, alkali-catalyzed transesterification reaction is much faster than Abbreviations: AV, acid value; PFAD, palm fatty acid distillate; FA, fatty acid; FAME, fatty acid methyl ester. ⇑ Corresponding author. Tel.: +82 2 2220 0488; fax: +82 2 2220 4007. E-mail address: [email protected] (Y.-K. Yeo). 0016-2361/$ - see front matter Ó 2011 Elsevier Ltd. All rights reserved. doi:10.1016/j.fuel.2011.08.063
acid-catalyzed transesterification and is popular in commercial production [9,10]. But alkali-catalyzed transesterification is suitable only for biodiesel production from feedstock containing low level of free fatty acid (FFA) such as refined vegetable oils. For efficient alkaline transesterification reactions, the feedstock should contain no more than 1 wt% FFA [11]. If the FFA level is too high, the soap generated during the transesterification inhibits separation of glycerol from the methyl esters after the reaction resulting in emulsion formation during the water wash and in induction of significant loss and poor quality of FAME product in the products [12]. Therefore, some low quality feedstock which is unrefined and much cheaper than the refined oil such as used cooking oils (2–7% FFA), animal fats (5–30%), palm fatty acid distillate (PFAD, 85–95%) and trap grease (100%) is definitely unavailable for alkaline transesterification processes. Several two-steps processes for the low-cost feedstock containing high content of free fatty acid (FFA) were then proposed [13–19]. In these processes, the first step is reduction of FFA content in the feedstock by esterification with methanol and acid catalyst such as sulfuric acid as a pre-treatment. The second step consists of transesterification process in which triglyceride portion of the feedstock reacts with methanol and base catalysts, usually sodium or potassium hydroxide, to form ester and glycerol. A major disadvantage of homogeneous catalyzed esterification reaction involving strong acid such as sulfuric acid is the difficulty in catalyst recovery and treatment, which generates a large amount of
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waste water, increasing the overall operation cost of the process. Furthermore, because the esterification process requires the equipment to be made of high-priced anti-corrosive material and multiple reaction stages to convert FFA to FAME sufficiently so that the product is suitable for the transesterification (FFA < 1%), the capital cost for the process is estimated not as competitive as traditional transesterification processes using refined oils. Recently, heterogeneous acid catalysts have been more widely favored over homogeneous ones since they are more separable and thus easier to recover. Especially there have been various studies on independent esterification processes without requiring transesterification based on the heterogeneous catalysts [20–27]. However, processes employing heterogeneous catalysts have not been widely commercialized. Competitiveness of those processes may be improved by increase of stability of catalysts during longterm operation and by application of efficient treatment methods for wastes generated during the catalyst-regenerating process. There have been only a few studies on non-catalytic esterification and/or transesterification reactions which lead to much simpler purification and environmentally friendly processes [28–33]. Diasakou et al. [28] investigated the non-catalytic thermal transesterification of soybean oil with methanol and performed study on reaction kinetics. Kusdiana and Saka [29] reported the effects of molar ratio of reactants and reaction temperature on the catalyst-free transesterification of rapeseed oil and proposed a simple method for the analysis of reaction kinetics. Yujaroen et al. [30] investigated the effects of temperature, molar ratio of methanol to fatty acid components and water content in the feed on the esterification of palm fatty acid distillate (PFAD). Most of these studies were conducted under pressurized conditions, i.e., supercritical or subcritical conditions of methanol. However the processes revealed in these works are not easily applicable to actual production of biodiesel due to significantly high production and capital costs required. In addition, these processes require severe operating conditions such as high pressure, high temperature and high molar ratio of methanol including uncertain safety aspects. On the other hand, Yamazaki et al. [31] studied the non-catalytic alcoholysis of sunflower oil for biodiesel production in a semi-batch bubble column reactor under atmospheric pressure. They analyzed the effects of reaction temperature, methanol feed flow rate, pressure, stirring rate and initial feed volume on the transesterification reaction. But, results of any in-depth studies which quantify the effect of operating conditions on the noncatalytic atmospheric alcoholysis were not shown in their work. Joelianingsih et al. [32] performed a kinetic study on the non-catalytic transesterification of palm oil in a semi-batch bubble column reactor at atmospheric pressure. They also investigated the effects of reaction temperatures on the rate constant, conversion, yield of methyl esters (ME) and composition of the reaction product. In another work by Joelianingsih et al. [12], the non-catalytic methyl esterification of oleic acid in a bubble column reactor was investigated and a kinetic study on the esterification was also revealed. However, the atmospheric transesterification and esterification investigated in these studies have a serious problem of significant decrease in yield of methyl ester only a small portion of which is recovered into reaction product. The yields of methyl ester revealed in the transesterification [32] and the esterification [12] are no more than 27.43% and 48.50%, respectively. Also, the conversions in these works only amount to 55.07% and 58.50%. It seems that the reason for such a low reactivity is a low solubility of methanol in the liquid phase at relatively high temperature (200–290 °C) and atmospheric pressure. The low yield and the conversion of the atmospheric reactions revealed in the works make it impossible for the processes to be applied to actual biodiesel production. In the present work, a novel method for the non-catalytic esterification of fatty acids, especially of palm fatty acid distillate (PFAD), is developed. In this method, the acid value of the product
is reduced to below 0.5 (mg KOH/g) which has not been reached by previous works within 180 min just in a single step reaction. For this reason the proposed method is readily applicable to real biodiesel production. By keeping the reaction condition at an optimal water-free range, i.e. at high enough to be above boiling point of water and also at below bubble points of methyl esters, not only reaction is completed beyond equilibrium but also no loss of methyl ester yield is achieved. The effects of temperature, pressure and methanol feed rate (i.e., molar ratio of methanol to fatty acids) on reactivity in a semi-batch CSTR (Continuously Stirred Tank Reactor) were also investigated. From the kinetic study on the non-catalytic esterification of PFAD, values of activation energy and frequency factor for the reaction were determined as well. Meanwhile, PFAD, used as raw material of the esterification reaction in this work, is a byproduct being inevitably generated in purification process of palm oil refinery, and so the price of PFAD is much cheaper than other refined oils which are major feed stocks for most of current biodiesel plants. Also, because PFAD consists of 85–95% fatty acids and 5–15% triglycerides both of which are available for biodiesel production, recently there have been a few trials that directly utilize PFAD as a feedstock for biodiesel production [27,30,33], but none of which have possibility of leading to actual production in commercial scale. The process for the non-catalytic single-step esterification of PFAD proposed in this work is readily applicable to actual biodiesel production and can be one of the most competitive processes due to its simplicity, excellent reaction yield and use of low-priced feedstock. 2. Experimental 2.1. Materials and apparatus The PFAD feedstock used in experiments is originated from Malaysia and is purchased from Sandakan Edible Oils SDN, BHD. The fatty acid (FA) content of the PFAD is 87.3% (Palmitic 41.6%, Oleic 33.5%, Linoleic 6.7%, Stearic 3.8%, Myristic 1.0%, Eicosanoic 0.4%, Eicosenoic 0.1%, Palmitoleic 0.2%) and the rest of the PFAD includes triglycerides 4.5%, diglycerides 3.7%, monoglycerides 2.6% and small amount of unknown impurities. The acid value (AV) of the PFAD sample used in experiments is 191.37 mg KOH/g. The purity of methanol which is supplied from SK Chemicals Ltd. exceeds 99.5%. Fig. 1 shows the experimental apparatus employed in this work. The reactor (1) is designed to perform semi-batch reactions. After the reactor is charged with reactants, methanol is fed into the bottom of the reactor via the methanol feed pump (2). Bubbles in the reactor are diminished and dispersed into the liquid phase by the impeller (3). The reactor is heated by the external jacket heater (4) and the internal temperature and RPM of the impeller are maintained as constant values by the controller (5). The excess methanol, water produced during the reaction and small amount of FA and FAME components vaporized and dispersed with water are condensed in the condenser (7) and collected in the receiver (8). Sampling is performed through the sampling line (13) with constant sampling time. The sample is transported to cooler (14) and cooled rapidly well below to 50 °C in order to prevent loss of low boiling components such as methanol and water by flashing. 2.2. Reaction procedure and conditions The PFAD feed used in the present work is in solid state at room temperature. Thus the feed is heated over 50 °C to be liquefied. The reactor is pressurized up to reaction pressure by injection of nitrogen gas after specific amount of the feed (860 g) is charged into the reactor. Agitation begins when the temperature reaches 60 °C
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H.J. Cho et al. / Fuel 93 (2012) 373–380 Table 1 Reaction conditions for the non-catalytic esterification employed in this work. No.
Temperature (°C)
Pressure (MPa)
Methanol feed rate (g/min)
1 2 3 4 5 6 7 8 9
230 250 270 290 290 290 290 290 290
0.85 0.85 0.85 0.85 0.10 0.50 1.20 0.85 0.85
2.4 2.4 2.4 2.4 2.4 2.4 2.4 1.2 3.6
Table 2 Reagents and apparatus used in this study.
Fig. 1. Schematic diagram of the reaction system for non-catalytic esterification. 1. Reactor (internal volume: 2L); 2. Methanol feed pump; 3. Impeller; 4. Electrical heater jacket; 5. Temperature and RPM controller; 6. Motor; 7. Condenser; 8. Receiver; 9. Pressure regulator; 10. Nitrogen charging line; 11. Thermocouple; 12. Pressure gauge; 13. Sampling line, 14. Cooler (double jacketed); 15. Chiller.
followed by subsequent heating up to reaction temperature. When the temperature of the reactor rises to the specific reaction temperature, methanol is fed into the reactor. The reactor temperature and speed of the impeller are maintained constant all the way during the reaction. The samples taken with constant sampling time are cooled below to 50 °C by the cooler. The reaction is terminated when the target reaction time is reached. Water produced during the reaction is vaporized and removed from the reactor so that the reaction rate is improved and the reaction proceeds to positive direction over equilibrium. It is important for FAME and FA not to be vaporized in order to minimize unnecessary loss. Fig. 2 shows the vapor pressures of water, methyl palmitate, methyl oleate, palmitic acid and oleic acid which are typical FAME and FA components with respect to temperature. As can be seen, most of water exists in vapor phase while FAMEs and FAs are hardly vaporized in present experimental conditions. In this work reaction behavior due to the changes in reaction temperature, pressure and methanol feed rates was analyzed and experi-
Heptadecanoicacid (Internal standard of fatty acid: >99.5%, Tokyo Chemical Industry Co.,Ltd.) N,O-Bis(trimethylsily)trifluoroacetamide(BSTFA) (Reagent grade, Sigma– Aldrich Co.) n-Heptane (Analytical grade: >99.9%, SK Chemicals Ltd.) Methylene chloride (Analytical grade: >99.9%, SK Chemicals Ltd.) GC system (Agilent 6890 N with FID detector) Column (Agilent, 122–5731, DB-5ht 30 m)
ments at atmospheric pressure were performed for the purpose of comparison with previous results on non-catalytic esterification study [12]. Table 1 summarizes reaction conditions used in the present work. 2.3. Analysis of FA and methanol contents FA content in samples is determined using the gas chromatography by mixing internal standard materials with samples. Typical reagents and apparatus used in the analysis of samples are listed in Table 2. The analysis method used to determine concentration of FA is based on internationally approved EN14103:2003 (Fat and oil derivatives – Fatty acid methyl esters (FAME) – Determination of ester and linolenic acid methyl ester contents) and modified newly in this work. The method starts with dissolution of heptadecanoic acid in n-heptane to give 100 mg/ml of the internal standard solution which is injected into the vial charged with 100 mg of sample by the amount of 200 lL through a syringe. Then 100 lL of N,OBis(trimethylsilyl)trifluoroacetamide (BSTFA) is added followed by silylation for 10 min in the oven kept at 60 °C. The FA content of the feed(C) is measured by the gas chromatography. The results represented in% (w/w) are calculated by the following equation:
C¼
Fig. 2. Vapor pressure of each componentand reaction conditionsused in this work (Methyl palmitate: calculated by Antoine equation [34]; others: calculated by AspenPlusÒ).
P ð AÞ AFI C FI V FI 100% AFI m
ð1Þ
P Here, A is sum of peak areas of various FA components, AFI is peak area of heptadecanoic acid, CFI is concentration of heptadecanoic acid solution (mg/ml), VFI is volume of heptadecanoic acid solution fed into the sample (ml), and m is mass of the sample (mg). As the gas chromatography column the non-polar column (Agilent, 122– 5731, DB-5ht 30 m) was chosen which is resistant to high temperature and is suitable for analysis of FA content. Determination of the concentration of methanol in a sample follows procedures and methods specified in EN 14110:2003 (Fat and oil derivatives – Fatty acid methyl esters (FAME) – Determination of methanol content). Water content was determined by using Karl Fisher analyzer following ISO 12937 (Determination of water – Coulometric Karl Fischer titration method). The acid value was measured according to the procedures and methods described in
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EN 14104:2003 (Fat and oil derivatives – Fatty acid methyl esters (FAME) – Determination of Acid value). 3. Results and discussion 3.1. Conversion In order to estimate the extent of reaction at the experimental conditions shown in Table 1, we employed the conversion of FA within PFAD given in Eq. (2) as an indicating index.
Xt ¼
ðC FA;0 C FA;t Þ 100% C FA
ð2Þ
Here, Xt is conversion of FA at time t and CFA,t, CFA,0 are concentrations of FA at reaction time t and initial time, respectively (mol/ ml). Under the reaction conditions considered, FAME is expected to be produced not only by esterification of FA but also by transesterification of triglyceride [31,32]. Thus the conversion was calculated based on the concentration of FA so that the extent of esterification of FA be estimated accurately. As shown in Fig. 3, the conversions calculated by Eq. (2) is inversely proportional to measured acid values, which means that the index we have chosen represents the extent of reaction adequately and validates the effectiveness of the new analysis method designed in the present work for FA component. 3.2. Effects of reaction temperature With the reaction pressure and methanol feed rate being kept constant as 0.85 MPa and 2.4 g/min, respectively, the conversions and changes in acid value according to reaction temperature are shown in Fig. 4. The region representing 90–180 min range in Fig. 4b is enlarged as shown in Fig. 4c. The reaction temperature adopted in this work is much higher than the boiling point of water (;173 °C at 0.85 MPa) produced during the reaction. Therefore, water is removed effectively from the reaction system and consequently backward reaction is inhibited. In short, we can accomplish very fast forward reaction even without using any catalysts. As shown in Fig. 4c, at the reaction temperature above 250 °C, the reaction is completed in 180 min and we can see that the resultant biodiesel products satisfy required criterion (AV < 0.5). It is observed that the acid value is not decreased down to 0.5 at 230 °C even in 240 min. It is believed that the acid value is lowered below 0.5 when 99.6% of FA in the feed is converted. So far, we could not find any published results that achieved desired target (AV < 0.5) using only single stage esterification reaction for 120–180 min with or without catalysts. Even under supercritical or subcritical conditions [27,30], the AV criterion of biodiesel was not achieved by single-stage reaction using strong acid catalyst such as sulfuric
Fig. 4. Effects of temperature on reaction (a) conversion of FA (b), (c) acid values measured (230–290 °C, 0.85 MPa, methanol feed rate: 2.4 g/min).
acid [33] or heterogeneous catalyst [27]. The main reason seems to be the lack of effective elimination of water from the reaction system. Comparing to other reaction methods developed so far, the proposed reaction system consisting of single-stage without catalyst can easily be commercialized and is promising due to its cost effective characteristics.
3.3. Effects of the reaction pressure
Fig. 3. A correlation of the reaction conversions with measured acid values. (290 °C, 0.85 MPa, methanol flow rate: 2.4 g/min).
Fig. 5a shows changes of AV of reaction products according to reaction pressure at constant temperature of 290 °C and constant methanol flow rate of 2.4 g/min. It can be seen that the AV decreases down to below 0.5 (target AV) in 240 min when the reaction pressure is greater than 0.50 MPa. The reaction is completed in 180 min especially when the reaction pressure is 0.85 MPa and 1.20 MPa. During the initial reaction stage (<90 min) the reaction rate increases as the reaction pressure decreases. But, as the reaction proceeds (90–240 min), we can see that the reaction rate increases with the reaction pressure. Fig. 5b shows the dependency of AV on reaction pressures. When the pressure is kept at 0.50 MPa, the reaction rate is retarded and it takes rather long time for AV to reach the target value. The AV of 0.92 mg KOH/g at
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H.J. Cho et al. / Fuel 93 (2012) 373–380 Table 3 Normal boiling products.
points
of
Component
B.P. (°C, at 101.3 kPa)
Oleic acid Methyl oleate Water Methanol
360.0 [35] 349.1 [36]
reactants/
100 64.2
Changes in reaction rate according to reaction pressure and differences in initial and final reaction rates seem to be caused by changes in concentration of methanol and water in the liquid phase. In the initial reaction stage the reaction proceeds intensely. Thus elimination of water from the reaction system is more effective than increase of the solubility of methanol to enhance the reaction rate (i.e., low reaction pressure is favored). But, in the late period of the reaction, the amount of water generated is smaller than that in the initial reaction stage and enhancement of methanol solubility is more advantageous to propel the reaction. Variations in concentrations of water and methanol in the liquid phase with respect to reaction time are shown in Fig. 5c and d, respectively. Rapid decrease in water concentration at the initial reaction stage (<90 min) demonstrates effective elimination of water generated during the reaction. We can see that large differences in water concentrations according to reaction pressure at the initial reaction stage shrink as the reaction proceeds. But, as can be seen in Fig. 5d, the concentration of methanol stays at some constant level after initial reaction stage. Moreover, the differences in the concentrations of methanol according to the reaction pressure do not change throughout the reaction. In fact, excess methanol fed into the reactor at the initial reaction stage is consumed rapidly due to the fast reaction rate. After the initial reaction stage, the concentration of methanol does not exhibit large variations because of the solubility which is proportional to the reaction pressure. Therefore we can say that the water content in the liquid phase governs the initial reaction rate while the methanol concentration determines the final reaction rate. This observation explains differences in reaction rates according to the reaction pressure at initial and final reaction stages as shown in Fig. 5a and b. The reaction pressure is closely related to the reaction yield. Lower reaction pressure causes decrease in the reaction yield due to the increase of the amount of loss caused by vaporization of FA and FAME. It is obvious that the mole fraction of FA and FAME in the vapor phase increases at low pressure. Table 3 and 4 show normal boiling points of typical FA and FAME components and variations in the product yield according to the reaction pressure, respectively. The yield greater than 100% is caused by the increase of molecular weight as water is substituted by the same amount (mol) of methanol during the esterification reaction. The average molecular weight of fatty acid component in PFAD used in experiments is 270.5 and that of methyl ester is 284.5. The maximum theoretical Fig. 5. Effects of pressure on reaction (a), (b) acid values (c) concentrations of water (d) concentrations of methanol (290 °C, methanol feed rate: 2.4 g/min).
90 min drops down to 0.5 mg KOH/g only after 240 min. But, when the pressure is kept at 1.20 MPa, the reaction proceeds faster and the AV of 2.1 mg KOH/g at 90 min decreases to 0.3 mg KOH/g at 180 min. Based on this observation we can construct more efficient esterification process by adopting two-step operation which consists of initial low pressure operation followed by high pressure operation when the reaction temperature is 290 °C and the methanol flow rate is 2.4 g/min.
Table 4 Effects of reaction pressure on the reaction yield. Pressure (MPa)
Product yielda (w/w (%))
0.10 0.50 0.85 1.20
94.4 100.0 102.8 103.2
a Product yield = (mass of reaction product)/ (mass of initial PFAD).
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H.J. Cho et al. / Fuel 93 (2012) 373–380
yield of the product generated from the esterification reaction of PFAD with methanol is 104.5%. As can be seen from Table 4, the yield is very low (94.4%) at atmospheric pressure due primarily to the increase of the loss by vaporization. On the other hand, the amount of the loss of FA in the form of vapor effluent with water and methanol during the reaction is very small at the reaction pressure greater than 0.50 MPa, considering the sampling loss and long reaction time (4 h). It is clear that atmospheric operation at relatively high reaction temperature is not adequate and that pressurized operation is required to prevent loss in reaction yield in the production of biodiesel. In short, the reaction pressure is the key variable affecting the reaction rate as well as the reaction yield. Therefore, for the commercial biodiesel production, the operating conditions tested in the present study may be preferred to the previous research on the non-catalytic esterification reaction performed under atmospheric pressure and high temperature [12].
3.5. Kinetics of the non-catalytic esterification in the semi-batch reaction system The non-catalytic esterification reaction performed in the semibatch reaction system can be regarded as a homogeneous reversible reaction. Assuming the 1st order reaction and considering mass balances for FA, we can represent the reaction rate as
dC FA ¼ kf C FA C MeOH kr C FAME C H2 O dt
ð3Þ
Here, CFA, CFAME, CMeOH, C H2 O are concentrations of fatty acid (FA), fatty acid methyl ester (FAME), methanol, and water, respectively (mol/ml) and kf, kr are the rate constants for forward reaction and for reverse reaction (ml/mol-min). Reaction conditions adopted in the present study aims continuous removal of water from the reaction system. Therefore we can ignore reverse reaction and Eq. (3) can be reduced to Eq. (4).
dC FA ¼ kf C FA C MeOH dt
3.4. Effects of the methanol feed rate
Fig. 6 shows changes of reaction conversion according to the methanol feed rate at constant temperature (290 °C) and pressure (0.85 MPa). As the methanol feed rate increases, the reaction rate grows fast. In fact, the reaction is completed within 180 min when the methanol rate is 2.4 g/min. But, for the flow rate of 1.2 g/min, we could not get satisfactory products even after 180 min (the AV is 3.0 mg KOH/g and the conversion is 98.4%). It is obvious that the methanol feed rate affects on the ratio of concentration of methanol to that of water in the liquid phase as in the case of the reaction pressure. Thus, a certain amount of methanol must be supplied to complete the reaction. As the amount of methanol increases, the water is removed more rapidly and effectively resulting in favorable reaction rates even though the solubility of the methanol is governed by temperature and pressure. But increase of the methanol feed rate causes higher utility cost as well as methanol recovery cost to make the process unattractive economically. It is important to identify the optimal methanol feed rate by considering reaction rate and manufacturing cost. Meanwhile recovery of methanol from the mixture of water and methanol flowing out from the reactor in the present study is considered to be much cost effective compared to those of conventional alkali-catalyzed biodiesel manufacturing processes. The vaporous methanol/water mixture flowing out from the reactor may be fed into the subsequent distillation process directly and the reboiler duty is reduced as much as the amount of the vaporization heat of the mixture feed. In conventional processes, methanol is recovered through vaporization of liquid methanol/water/ glycerol mixture flowing out from decanting process. For this reason, most of conventional processes require more energy in recovering methanol compared to the reaction scheme in the present work.
Moreover, since excess methanol is fed into the reaction system as in the case of Joelianingsih et al. [12], CMeOH can be regarded as constant and Eq. (4) may be rewritten in terms of CFA as
ð4Þ
dC FA 0 ¼ kf C FA dt
ð5Þ 0
Then the reduced rate constant kf of the forward reaction is given by the following equation:
ln
C FA;t C FA;0
0
¼ kf t
ð6Þ
Table 5 summarizes rate constants at different reaction temperatures obtained by plugging the results of experiments carried out at constant pressure (0.85 MPa) and methanol feed rate (2.4 g/min) while varying reaction temperature into Eq. (6). Results of Joelianingsih et al. [12] are also shown for comparison. From Table 5, 0 we can see that rate constants kf by present study are about 4 times greater than those obtained by Joelianingsih et al. [12]. High rate constants by present study seem to be caused by the large methanol concentration at 0.85 MPa, which is about 20 times greater than that measured at atmospheric pressure. Concentrations of FA and FAME computed by substituting rate constants in Table 5 into Eq. (5) are shown in Fig. 7 with measured values. Table 6 shows values of root-mean squares errors (RMSE) denoting discrepancy between kinetic equations and experimental data. As can be seen, computational results agree well with experimental results. Therefore, rate constants obtained from the present study and the reduced model given by Eq. (5) have some guiding significance to describe the reaction mechanism. 0 The rate constant kf satisfies Arrhenius Eq. (7). 0
ln kf ¼ ln A
Ea RT
ð7Þ 0
Fig. 8 displays relationship between temperature and kf given in Table 5. We can see that the results are well fitted by Arrhenius equation. The activation energy is 17.74 kJ/mol and the frequency Table 5 Comparison of rate constants (present study: 0.85 MPa, Joelianingsih et al. [12]: 101.3 kPa). Present study
Fig. 6. Effects of methanol feed rate on the reaction conversion (290 °C, 0.85 MPa).
Joelianingsih et al. [12]
T (°C)
0 kf
230 250 270 290
0.0299 0.0375 0.0405 0.0481
1
(min
)
R2
T (°C)
kf (min1)
0.954 0.974 0.966 0.961
200 225 250
0.0052 0.0073 0.0095
0
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H.J. Cho et al. / Fuel 93 (2012) 373–380 Table 6 Values of RMSE for each component (103). Temperature(°C)
FA FAME Overall
230
250
270
290
0.1792 0.3474 0.2682
0.2787 0.2058 0.2376
0.2637 0.2046 0.2290
0.3197 0.2281 0.2694
Fig. 8. Arrhenius plot for the esterification of PFAD (0.85 MPa, methanol feed rate: 2.4 g/min).
while the activation energy by Joelianingsih et al. [12] (24.8 kJ/ mol) is a little larger than present value. This means that the non-catalytic reaction is accelerated as temperature increases. 4. Conclusions A single-step non-catalytic esterification reaction using PFAD feedstock is investigated in this study. The PFAD feedstock contains 85–95% of various fatty acid components and resulting acid value satisfies the biodiesel product criterion of 0.5 (mg KOH/g). Reaction conditions are selected so that water being generated during the reaction is removed rapidly (temperature: >250 °C, pressure: 0.85–1.20 MPa, reaction time: <180 min). Effects of reaction temperature, pressure and the methanol flow rate are analyzed and reaction rate constants are determined by the 1st-order model of FA concentration. The concentration profiles of FA and FAME computed using aforementioned rate constants and reduced 1st-order model fit well with experimental data. From rate constants and Arrhenius relation, the activation energy and the frequency factor were found to be 17.74 kJ/mol and 2.12 min1, respectively. The main characteristic of the proposed reaction method is the employment of single-step reaction without catalyst to yield commercial grade biodiesel product with little loss. For this reason, the present reaction technique can be readily commercialized. Moreover, even low grade waste fatty materials such as waste greases can be used as feedstock, which was not permitted in traditional transesterification reaction due to high fatty acid content. References Fig. 7. Concentrations of FA and FAME: comparison between computational results (based on the 1st-order model) and experimental results (a) 230 °C, (b) 250 °C, (c) 270 °C, and (d) 290 °C.
factor is 2.12 min1. From comparison with values by Joelianingsih et al. [12], it appears that the frequency factors are about the same
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[5] Demirbas A. Biodiesel: a realistic fuel alternative for diesel engines. 1st ed. London: Springer Verlag; 2008. [6] Sendzikiene E, Makareviciene V, Janulis P. Influence of fuel oxygen content on diesel engine exhaust emissions. Renew Energy 2006;31:2505–12. [7] Schumacher LG, Borgelt SC, Hires WG. Fueling a diesel engine with methylester soybean oil. Appl Eng Agric 1995;11(1):37–40. [8] Marchetti JM, Miguel VU, Errazu AF. Possible methods for biodiesel production. Renew Sustain Energy Rev 2007;11:1300–11. [9] Ma F, Hanna MA. Biodiesel production: a review. Bioresour Technol 1999;70:1–15. [10] Vincente G, Martinez M, Aracil J. Integrated biodiesel production: a comparison of different homogeneous catalysts systems. Bioresour Technol 2004;92:297–305. [11] Berrios M, Siles J, Martin MA, Martin AA. Kinetic study of the esterification of free fatty acids (FFA) in sunflower oil. Fuel 2007;86:2383–8. [12] Joelianingsih, Nabetani H, Hagiwara S, Sagara Y, Soerawidjaya TH, Tambunan AH, et al. Performance of a bubble column reactor for the non-catalytic methyl esterification of free fatty acids at atmospheric pressure. J Chem Eng Japan 2007;40(9):780–5. [13] Jeromin L, Peukert E, Wollmann G. Process for the pre-esterification of free fatty acids in fats and oils. US Patent 4698,186; 1987. [14] Ramadhas AS, Jayaraj S, Muraleedharan C. Biodiesel production from high FFA rubber seed oil. Fuel 2004;84:335–40. [15] Veljkovic VB, Lakicevic SH, Stamenkovic OS, Todorovi ZB, Lazic ML. Biodiesel production from tobacco seed oil with a high content of free fatty acids. Fuel 2006;85:2671–5. [16] Wang Y, Ou S, Liu P, Xue F, Tang S. Comparison of two different processes to synthesize biodiesel by waste cooking oil. J Molec Catal A: Chemicals 2006;252:107–12. [17] Shashikant VG, Raheman H. Biodiesel production from mahua oil having high free fatty acids. Biom Bioe 2005;28(6):601–5. [18] Shashikant VG, Raheman H. Process optimization for biodiesel production from mahua (Madhucaindica) oil using response surface methodology. Bioresour Technol 2006;97:379–84. [19] Zheng S, Kates M, Dube MA, McLean DD. Acid-catalyzed production of biodiesel from waste frying oil. Biom Bioe 2006;30:267–72. [20] Kiss AA, Omota F, Dimian AC, Rothenberg G. The heterogeneous advantage: biodiesel by catalytic reactive distillation. Top Catal 2006;40:141–50. [21] Jitputti J, Kitiyanan B, Rangsunvigit P, Bunyakait K, Attanatho L, Jenvanitpanjakul P. Transesterification of crude palm kernel oil and crude coconut oil by different solid catalysts. Chem Eng J 2006;116:61–6.
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A single step non-catalytic esterification of palm fatty acid distillate (PFAD) for biodiesel production Hyun Jun Cho a, Soo Hyun Kim b, Seok Won Hong a, Yeong-Koo Yeo a,⇑ a b
Department of Chemical Engineering, Hanyang University, Haengdang-dong, Sungdong-gu, Seoul, Republic of Korea Chemicals R&D Center, SK Chemicals, 686 Sampyeong-dong, Bundang-gu, Seongnam-si, Gyeonggi-do, Republic of Korea
a r t i c l e
i n f o
Article history: Received 11 October 2010 Received in revised form 26 August 2011 Accepted 26 August 2011 Available online 13 September 2011 Keywords: Biodiesel Non-catalytic esterification Palm fatty acid distillate Water-free reaction
a b s t r a c t In this work, the single step method for non-catalytic esterification of palm fatty acid distillate (PFAD), which is readily applicable to actual production of biodiesel, was investigated. In this method, the esterification reaction is accomplished in a single step by ensuring water-free reaction conditions and the acid value is reduced to below 0.5 (mg KOH/g) which has not been possible in previous methods. The reaction was completed (<0.5 mg KOH/g) within 180 min at relatively high temperature (>250 °C) enough to be above boiling point of water and at moderate pressure (0.85–1.20 MPa) without any catalyst. The effects of temperature, methanol feed rate and pressure on a semi-batch reaction were investigated and the optimal values of these variables were found (temperature: 290 °C, pressure: 0.85 MPa, feed rate: 2.4 g/min). The acid value was reduced from 191.4 to 0.36 (mg KOH/g) just in 180 min at these conditions. From the kinetic study on non-catalytic esterification of PFAD, it was found that the activation energy is 17.74 kJ/ mol and the frequency factor is 2.12 min1. Ó 2011 Elsevier Ltd. All rights reserved.
1. Introduction As an alternative diesel fuel, biodiesel is made out of renewable biomass such as vegetable oils, animal fats, and waste grease. It is well known that the utilization of biodiesel is carbon–neutral and contributable to the reduction of CO2 emission [1,2]. Due to its biodegradability, nontoxic property and low emission profiles, biodiesel has attracted lots of attention and has been considered environmentally beneficial [3,4]. Because biodiesel contains no sulfur but about 11% of oxygen (w/w), we can expect SOx free emission and a significant reduction of pollutant emissions such as unburned hydrocarbons, carbon monoxide and particulate matters compared with petroleum-derived diesel fuel [5,6]. In addition to the environmental benefits, primary advantage of biodiesel is that none- or few modifications of current diesel engine system are needed to adopt biodiesel blends or even pure biodiesel [7]. In general, biodiesel, fatty acid methyl ester (FAME), is efficiently produced from the transesterification of vegetable oil or animal fats or from esterification of fatty acids with short chain alcohols in the presence of homogeneous or heterogeneous alkali- and/or acid–based catalysts [8]. Among catalytic processes, alkali-catalyzed transesterification reaction is much faster than Abbreviations: AV, acid value; PFAD, palm fatty acid distillate; FA, fatty acid; FAME, fatty acid methyl ester. ⇑ Corresponding author. Tel.: +82 2 2220 0488; fax: +82 2 2220 4007. E-mail address: [email protected] (Y.-K. Yeo). 0016-2361/$ - see front matter Ó 2011 Elsevier Ltd. All rights reserved. doi:10.1016/j.fuel.2011.08.063
acid-catalyzed transesterification and is popular in commercial production [9,10]. But alkali-catalyzed transesterification is suitable only for biodiesel production from feedstock containing low level of free fatty acid (FFA) such as refined vegetable oils. For efficient alkaline transesterification reactions, the feedstock should contain no more than 1 wt% FFA [11]. If the FFA level is too high, the soap generated during the transesterification inhibits separation of glycerol from the methyl esters after the reaction resulting in emulsion formation during the water wash and in induction of significant loss and poor quality of FAME product in the products [12]. Therefore, some low quality feedstock which is unrefined and much cheaper than the refined oil such as used cooking oils (2–7% FFA), animal fats (5–30%), palm fatty acid distillate (PFAD, 85–95%) and trap grease (100%) is definitely unavailable for alkaline transesterification processes. Several two-steps processes for the low-cost feedstock containing high content of free fatty acid (FFA) were then proposed [13–19]. In these processes, the first step is reduction of FFA content in the feedstock by esterification with methanol and acid catalyst such as sulfuric acid as a pre-treatment. The second step consists of transesterification process in which triglyceride portion of the feedstock reacts with methanol and base catalysts, usually sodium or potassium hydroxide, to form ester and glycerol. A major disadvantage of homogeneous catalyzed esterification reaction involving strong acid such as sulfuric acid is the difficulty in catalyst recovery and treatment, which generates a large amount of
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waste water, increasing the overall operation cost of the process. Furthermore, because the esterification process requires the equipment to be made of high-priced anti-corrosive material and multiple reaction stages to convert FFA to FAME sufficiently so that the product is suitable for the transesterification (FFA < 1%), the capital cost for the process is estimated not as competitive as traditional transesterification processes using refined oils. Recently, heterogeneous acid catalysts have been more widely favored over homogeneous ones since they are more separable and thus easier to recover. Especially there have been various studies on independent esterification processes without requiring transesterification based on the heterogeneous catalysts [20–27]. However, processes employing heterogeneous catalysts have not been widely commercialized. Competitiveness of those processes may be improved by increase of stability of catalysts during longterm operation and by application of efficient treatment methods for wastes generated during the catalyst-regenerating process. There have been only a few studies on non-catalytic esterification and/or transesterification reactions which lead to much simpler purification and environmentally friendly processes [28–33]. Diasakou et al. [28] investigated the non-catalytic thermal transesterification of soybean oil with methanol and performed study on reaction kinetics. Kusdiana and Saka [29] reported the effects of molar ratio of reactants and reaction temperature on the catalyst-free transesterification of rapeseed oil and proposed a simple method for the analysis of reaction kinetics. Yujaroen et al. [30] investigated the effects of temperature, molar ratio of methanol to fatty acid components and water content in the feed on the esterification of palm fatty acid distillate (PFAD). Most of these studies were conducted under pressurized conditions, i.e., supercritical or subcritical conditions of methanol. However the processes revealed in these works are not easily applicable to actual production of biodiesel due to significantly high production and capital costs required. In addition, these processes require severe operating conditions such as high pressure, high temperature and high molar ratio of methanol including uncertain safety aspects. On the other hand, Yamazaki et al. [31] studied the non-catalytic alcoholysis of sunflower oil for biodiesel production in a semi-batch bubble column reactor under atmospheric pressure. They analyzed the effects of reaction temperature, methanol feed flow rate, pressure, stirring rate and initial feed volume on the transesterification reaction. But, results of any in-depth studies which quantify the effect of operating conditions on the noncatalytic atmospheric alcoholysis were not shown in their work. Joelianingsih et al. [32] performed a kinetic study on the non-catalytic transesterification of palm oil in a semi-batch bubble column reactor at atmospheric pressure. They also investigated the effects of reaction temperatures on the rate constant, conversion, yield of methyl esters (ME) and composition of the reaction product. In another work by Joelianingsih et al. [12], the non-catalytic methyl esterification of oleic acid in a bubble column reactor was investigated and a kinetic study on the esterification was also revealed. However, the atmospheric transesterification and esterification investigated in these studies have a serious problem of significant decrease in yield of methyl ester only a small portion of which is recovered into reaction product. The yields of methyl ester revealed in the transesterification [32] and the esterification [12] are no more than 27.43% and 48.50%, respectively. Also, the conversions in these works only amount to 55.07% and 58.50%. It seems that the reason for such a low reactivity is a low solubility of methanol in the liquid phase at relatively high temperature (200–290 °C) and atmospheric pressure. The low yield and the conversion of the atmospheric reactions revealed in the works make it impossible for the processes to be applied to actual biodiesel production. In the present work, a novel method for the non-catalytic esterification of fatty acids, especially of palm fatty acid distillate (PFAD), is developed. In this method, the acid value of the product
is reduced to below 0.5 (mg KOH/g) which has not been reached by previous works within 180 min just in a single step reaction. For this reason the proposed method is readily applicable to real biodiesel production. By keeping the reaction condition at an optimal water-free range, i.e. at high enough to be above boiling point of water and also at below bubble points of methyl esters, not only reaction is completed beyond equilibrium but also no loss of methyl ester yield is achieved. The effects of temperature, pressure and methanol feed rate (i.e., molar ratio of methanol to fatty acids) on reactivity in a semi-batch CSTR (Continuously Stirred Tank Reactor) were also investigated. From the kinetic study on the non-catalytic esterification of PFAD, values of activation energy and frequency factor for the reaction were determined as well. Meanwhile, PFAD, used as raw material of the esterification reaction in this work, is a byproduct being inevitably generated in purification process of palm oil refinery, and so the price of PFAD is much cheaper than other refined oils which are major feed stocks for most of current biodiesel plants. Also, because PFAD consists of 85–95% fatty acids and 5–15% triglycerides both of which are available for biodiesel production, recently there have been a few trials that directly utilize PFAD as a feedstock for biodiesel production [27,30,33], but none of which have possibility of leading to actual production in commercial scale. The process for the non-catalytic single-step esterification of PFAD proposed in this work is readily applicable to actual biodiesel production and can be one of the most competitive processes due to its simplicity, excellent reaction yield and use of low-priced feedstock. 2. Experimental 2.1. Materials and apparatus The PFAD feedstock used in experiments is originated from Malaysia and is purchased from Sandakan Edible Oils SDN, BHD. The fatty acid (FA) content of the PFAD is 87.3% (Palmitic 41.6%, Oleic 33.5%, Linoleic 6.7%, Stearic 3.8%, Myristic 1.0%, Eicosanoic 0.4%, Eicosenoic 0.1%, Palmitoleic 0.2%) and the rest of the PFAD includes triglycerides 4.5%, diglycerides 3.7%, monoglycerides 2.6% and small amount of unknown impurities. The acid value (AV) of the PFAD sample used in experiments is 191.37 mg KOH/g. The purity of methanol which is supplied from SK Chemicals Ltd. exceeds 99.5%. Fig. 1 shows the experimental apparatus employed in this work. The reactor (1) is designed to perform semi-batch reactions. After the reactor is charged with reactants, methanol is fed into the bottom of the reactor via the methanol feed pump (2). Bubbles in the reactor are diminished and dispersed into the liquid phase by the impeller (3). The reactor is heated by the external jacket heater (4) and the internal temperature and RPM of the impeller are maintained as constant values by the controller (5). The excess methanol, water produced during the reaction and small amount of FA and FAME components vaporized and dispersed with water are condensed in the condenser (7) and collected in the receiver (8). Sampling is performed through the sampling line (13) with constant sampling time. The sample is transported to cooler (14) and cooled rapidly well below to 50 °C in order to prevent loss of low boiling components such as methanol and water by flashing. 2.2. Reaction procedure and conditions The PFAD feed used in the present work is in solid state at room temperature. Thus the feed is heated over 50 °C to be liquefied. The reactor is pressurized up to reaction pressure by injection of nitrogen gas after specific amount of the feed (860 g) is charged into the reactor. Agitation begins when the temperature reaches 60 °C
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H.J. Cho et al. / Fuel 93 (2012) 373–380 Table 1 Reaction conditions for the non-catalytic esterification employed in this work. No.
Temperature (°C)
Pressure (MPa)
Methanol feed rate (g/min)
1 2 3 4 5 6 7 8 9
230 250 270 290 290 290 290 290 290
0.85 0.85 0.85 0.85 0.10 0.50 1.20 0.85 0.85
2.4 2.4 2.4 2.4 2.4 2.4 2.4 1.2 3.6
Table 2 Reagents and apparatus used in this study.
Fig. 1. Schematic diagram of the reaction system for non-catalytic esterification. 1. Reactor (internal volume: 2L); 2. Methanol feed pump; 3. Impeller; 4. Electrical heater jacket; 5. Temperature and RPM controller; 6. Motor; 7. Condenser; 8. Receiver; 9. Pressure regulator; 10. Nitrogen charging line; 11. Thermocouple; 12. Pressure gauge; 13. Sampling line, 14. Cooler (double jacketed); 15. Chiller.
followed by subsequent heating up to reaction temperature. When the temperature of the reactor rises to the specific reaction temperature, methanol is fed into the reactor. The reactor temperature and speed of the impeller are maintained constant all the way during the reaction. The samples taken with constant sampling time are cooled below to 50 °C by the cooler. The reaction is terminated when the target reaction time is reached. Water produced during the reaction is vaporized and removed from the reactor so that the reaction rate is improved and the reaction proceeds to positive direction over equilibrium. It is important for FAME and FA not to be vaporized in order to minimize unnecessary loss. Fig. 2 shows the vapor pressures of water, methyl palmitate, methyl oleate, palmitic acid and oleic acid which are typical FAME and FA components with respect to temperature. As can be seen, most of water exists in vapor phase while FAMEs and FAs are hardly vaporized in present experimental conditions. In this work reaction behavior due to the changes in reaction temperature, pressure and methanol feed rates was analyzed and experi-
Heptadecanoicacid (Internal standard of fatty acid: >99.5%, Tokyo Chemical Industry Co.,Ltd.) N,O-Bis(trimethylsily)trifluoroacetamide(BSTFA) (Reagent grade, Sigma– Aldrich Co.) n-Heptane (Analytical grade: >99.9%, SK Chemicals Ltd.) Methylene chloride (Analytical grade: >99.9%, SK Chemicals Ltd.) GC system (Agilent 6890 N with FID detector) Column (Agilent, 122–5731, DB-5ht 30 m)
ments at atmospheric pressure were performed for the purpose of comparison with previous results on non-catalytic esterification study [12]. Table 1 summarizes reaction conditions used in the present work. 2.3. Analysis of FA and methanol contents FA content in samples is determined using the gas chromatography by mixing internal standard materials with samples. Typical reagents and apparatus used in the analysis of samples are listed in Table 2. The analysis method used to determine concentration of FA is based on internationally approved EN14103:2003 (Fat and oil derivatives – Fatty acid methyl esters (FAME) – Determination of ester and linolenic acid methyl ester contents) and modified newly in this work. The method starts with dissolution of heptadecanoic acid in n-heptane to give 100 mg/ml of the internal standard solution which is injected into the vial charged with 100 mg of sample by the amount of 200 lL through a syringe. Then 100 lL of N,OBis(trimethylsilyl)trifluoroacetamide (BSTFA) is added followed by silylation for 10 min in the oven kept at 60 °C. The FA content of the feed(C) is measured by the gas chromatography. The results represented in% (w/w) are calculated by the following equation:
C¼
Fig. 2. Vapor pressure of each componentand reaction conditionsused in this work (Methyl palmitate: calculated by Antoine equation [34]; others: calculated by AspenPlusÒ).
P ð AÞ AFI C FI V FI 100% AFI m
ð1Þ
P Here, A is sum of peak areas of various FA components, AFI is peak area of heptadecanoic acid, CFI is concentration of heptadecanoic acid solution (mg/ml), VFI is volume of heptadecanoic acid solution fed into the sample (ml), and m is mass of the sample (mg). As the gas chromatography column the non-polar column (Agilent, 122– 5731, DB-5ht 30 m) was chosen which is resistant to high temperature and is suitable for analysis of FA content. Determination of the concentration of methanol in a sample follows procedures and methods specified in EN 14110:2003 (Fat and oil derivatives – Fatty acid methyl esters (FAME) – Determination of methanol content). Water content was determined by using Karl Fisher analyzer following ISO 12937 (Determination of water – Coulometric Karl Fischer titration method). The acid value was measured according to the procedures and methods described in
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EN 14104:2003 (Fat and oil derivatives – Fatty acid methyl esters (FAME) – Determination of Acid value). 3. Results and discussion 3.1. Conversion In order to estimate the extent of reaction at the experimental conditions shown in Table 1, we employed the conversion of FA within PFAD given in Eq. (2) as an indicating index.
Xt ¼
ðC FA;0 C FA;t Þ 100% C FA
ð2Þ
Here, Xt is conversion of FA at time t and CFA,t, CFA,0 are concentrations of FA at reaction time t and initial time, respectively (mol/ ml). Under the reaction conditions considered, FAME is expected to be produced not only by esterification of FA but also by transesterification of triglyceride [31,32]. Thus the conversion was calculated based on the concentration of FA so that the extent of esterification of FA be estimated accurately. As shown in Fig. 3, the conversions calculated by Eq. (2) is inversely proportional to measured acid values, which means that the index we have chosen represents the extent of reaction adequately and validates the effectiveness of the new analysis method designed in the present work for FA component. 3.2. Effects of reaction temperature With the reaction pressure and methanol feed rate being kept constant as 0.85 MPa and 2.4 g/min, respectively, the conversions and changes in acid value according to reaction temperature are shown in Fig. 4. The region representing 90–180 min range in Fig. 4b is enlarged as shown in Fig. 4c. The reaction temperature adopted in this work is much higher than the boiling point of water (;173 °C at 0.85 MPa) produced during the reaction. Therefore, water is removed effectively from the reaction system and consequently backward reaction is inhibited. In short, we can accomplish very fast forward reaction even without using any catalysts. As shown in Fig. 4c, at the reaction temperature above 250 °C, the reaction is completed in 180 min and we can see that the resultant biodiesel products satisfy required criterion (AV < 0.5). It is observed that the acid value is not decreased down to 0.5 at 230 °C even in 240 min. It is believed that the acid value is lowered below 0.5 when 99.6% of FA in the feed is converted. So far, we could not find any published results that achieved desired target (AV < 0.5) using only single stage esterification reaction for 120–180 min with or without catalysts. Even under supercritical or subcritical conditions [27,30], the AV criterion of biodiesel was not achieved by single-stage reaction using strong acid catalyst such as sulfuric
Fig. 4. Effects of temperature on reaction (a) conversion of FA (b), (c) acid values measured (230–290 °C, 0.85 MPa, methanol feed rate: 2.4 g/min).
acid [33] or heterogeneous catalyst [27]. The main reason seems to be the lack of effective elimination of water from the reaction system. Comparing to other reaction methods developed so far, the proposed reaction system consisting of single-stage without catalyst can easily be commercialized and is promising due to its cost effective characteristics.
3.3. Effects of the reaction pressure
Fig. 3. A correlation of the reaction conversions with measured acid values. (290 °C, 0.85 MPa, methanol flow rate: 2.4 g/min).
Fig. 5a shows changes of AV of reaction products according to reaction pressure at constant temperature of 290 °C and constant methanol flow rate of 2.4 g/min. It can be seen that the AV decreases down to below 0.5 (target AV) in 240 min when the reaction pressure is greater than 0.50 MPa. The reaction is completed in 180 min especially when the reaction pressure is 0.85 MPa and 1.20 MPa. During the initial reaction stage (<90 min) the reaction rate increases as the reaction pressure decreases. But, as the reaction proceeds (90–240 min), we can see that the reaction rate increases with the reaction pressure. Fig. 5b shows the dependency of AV on reaction pressures. When the pressure is kept at 0.50 MPa, the reaction rate is retarded and it takes rather long time for AV to reach the target value. The AV of 0.92 mg KOH/g at
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H.J. Cho et al. / Fuel 93 (2012) 373–380 Table 3 Normal boiling products.
points
of
Component
B.P. (°C, at 101.3 kPa)
Oleic acid Methyl oleate Water Methanol
360.0 [35] 349.1 [36]
reactants/
100 64.2
Changes in reaction rate according to reaction pressure and differences in initial and final reaction rates seem to be caused by changes in concentration of methanol and water in the liquid phase. In the initial reaction stage the reaction proceeds intensely. Thus elimination of water from the reaction system is more effective than increase of the solubility of methanol to enhance the reaction rate (i.e., low reaction pressure is favored). But, in the late period of the reaction, the amount of water generated is smaller than that in the initial reaction stage and enhancement of methanol solubility is more advantageous to propel the reaction. Variations in concentrations of water and methanol in the liquid phase with respect to reaction time are shown in Fig. 5c and d, respectively. Rapid decrease in water concentration at the initial reaction stage (<90 min) demonstrates effective elimination of water generated during the reaction. We can see that large differences in water concentrations according to reaction pressure at the initial reaction stage shrink as the reaction proceeds. But, as can be seen in Fig. 5d, the concentration of methanol stays at some constant level after initial reaction stage. Moreover, the differences in the concentrations of methanol according to the reaction pressure do not change throughout the reaction. In fact, excess methanol fed into the reactor at the initial reaction stage is consumed rapidly due to the fast reaction rate. After the initial reaction stage, the concentration of methanol does not exhibit large variations because of the solubility which is proportional to the reaction pressure. Therefore we can say that the water content in the liquid phase governs the initial reaction rate while the methanol concentration determines the final reaction rate. This observation explains differences in reaction rates according to the reaction pressure at initial and final reaction stages as shown in Fig. 5a and b. The reaction pressure is closely related to the reaction yield. Lower reaction pressure causes decrease in the reaction yield due to the increase of the amount of loss caused by vaporization of FA and FAME. It is obvious that the mole fraction of FA and FAME in the vapor phase increases at low pressure. Table 3 and 4 show normal boiling points of typical FA and FAME components and variations in the product yield according to the reaction pressure, respectively. The yield greater than 100% is caused by the increase of molecular weight as water is substituted by the same amount (mol) of methanol during the esterification reaction. The average molecular weight of fatty acid component in PFAD used in experiments is 270.5 and that of methyl ester is 284.5. The maximum theoretical Fig. 5. Effects of pressure on reaction (a), (b) acid values (c) concentrations of water (d) concentrations of methanol (290 °C, methanol feed rate: 2.4 g/min).
90 min drops down to 0.5 mg KOH/g only after 240 min. But, when the pressure is kept at 1.20 MPa, the reaction proceeds faster and the AV of 2.1 mg KOH/g at 90 min decreases to 0.3 mg KOH/g at 180 min. Based on this observation we can construct more efficient esterification process by adopting two-step operation which consists of initial low pressure operation followed by high pressure operation when the reaction temperature is 290 °C and the methanol flow rate is 2.4 g/min.
Table 4 Effects of reaction pressure on the reaction yield. Pressure (MPa)
Product yielda (w/w (%))
0.10 0.50 0.85 1.20
94.4 100.0 102.8 103.2
a Product yield = (mass of reaction product)/ (mass of initial PFAD).
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H.J. Cho et al. / Fuel 93 (2012) 373–380
yield of the product generated from the esterification reaction of PFAD with methanol is 104.5%. As can be seen from Table 4, the yield is very low (94.4%) at atmospheric pressure due primarily to the increase of the loss by vaporization. On the other hand, the amount of the loss of FA in the form of vapor effluent with water and methanol during the reaction is very small at the reaction pressure greater than 0.50 MPa, considering the sampling loss and long reaction time (4 h). It is clear that atmospheric operation at relatively high reaction temperature is not adequate and that pressurized operation is required to prevent loss in reaction yield in the production of biodiesel. In short, the reaction pressure is the key variable affecting the reaction rate as well as the reaction yield. Therefore, for the commercial biodiesel production, the operating conditions tested in the present study may be preferred to the previous research on the non-catalytic esterification reaction performed under atmospheric pressure and high temperature [12].
3.5. Kinetics of the non-catalytic esterification in the semi-batch reaction system The non-catalytic esterification reaction performed in the semibatch reaction system can be regarded as a homogeneous reversible reaction. Assuming the 1st order reaction and considering mass balances for FA, we can represent the reaction rate as
dC FA ¼ kf C FA C MeOH kr C FAME C H2 O dt
ð3Þ
Here, CFA, CFAME, CMeOH, C H2 O are concentrations of fatty acid (FA), fatty acid methyl ester (FAME), methanol, and water, respectively (mol/ml) and kf, kr are the rate constants for forward reaction and for reverse reaction (ml/mol-min). Reaction conditions adopted in the present study aims continuous removal of water from the reaction system. Therefore we can ignore reverse reaction and Eq. (3) can be reduced to Eq. (4).
dC FA ¼ kf C FA C MeOH dt
3.4. Effects of the methanol feed rate
Fig. 6 shows changes of reaction conversion according to the methanol feed rate at constant temperature (290 °C) and pressure (0.85 MPa). As the methanol feed rate increases, the reaction rate grows fast. In fact, the reaction is completed within 180 min when the methanol rate is 2.4 g/min. But, for the flow rate of 1.2 g/min, we could not get satisfactory products even after 180 min (the AV is 3.0 mg KOH/g and the conversion is 98.4%). It is obvious that the methanol feed rate affects on the ratio of concentration of methanol to that of water in the liquid phase as in the case of the reaction pressure. Thus, a certain amount of methanol must be supplied to complete the reaction. As the amount of methanol increases, the water is removed more rapidly and effectively resulting in favorable reaction rates even though the solubility of the methanol is governed by temperature and pressure. But increase of the methanol feed rate causes higher utility cost as well as methanol recovery cost to make the process unattractive economically. It is important to identify the optimal methanol feed rate by considering reaction rate and manufacturing cost. Meanwhile recovery of methanol from the mixture of water and methanol flowing out from the reactor in the present study is considered to be much cost effective compared to those of conventional alkali-catalyzed biodiesel manufacturing processes. The vaporous methanol/water mixture flowing out from the reactor may be fed into the subsequent distillation process directly and the reboiler duty is reduced as much as the amount of the vaporization heat of the mixture feed. In conventional processes, methanol is recovered through vaporization of liquid methanol/water/ glycerol mixture flowing out from decanting process. For this reason, most of conventional processes require more energy in recovering methanol compared to the reaction scheme in the present work.
Moreover, since excess methanol is fed into the reaction system as in the case of Joelianingsih et al. [12], CMeOH can be regarded as constant and Eq. (4) may be rewritten in terms of CFA as
ð4Þ
dC FA 0 ¼ kf C FA dt
ð5Þ 0
Then the reduced rate constant kf of the forward reaction is given by the following equation:
ln
C FA;t C FA;0
0
¼ kf t
ð6Þ
Table 5 summarizes rate constants at different reaction temperatures obtained by plugging the results of experiments carried out at constant pressure (0.85 MPa) and methanol feed rate (2.4 g/min) while varying reaction temperature into Eq. (6). Results of Joelianingsih et al. [12] are also shown for comparison. From Table 5, 0 we can see that rate constants kf by present study are about 4 times greater than those obtained by Joelianingsih et al. [12]. High rate constants by present study seem to be caused by the large methanol concentration at 0.85 MPa, which is about 20 times greater than that measured at atmospheric pressure. Concentrations of FA and FAME computed by substituting rate constants in Table 5 into Eq. (5) are shown in Fig. 7 with measured values. Table 6 shows values of root-mean squares errors (RMSE) denoting discrepancy between kinetic equations and experimental data. As can be seen, computational results agree well with experimental results. Therefore, rate constants obtained from the present study and the reduced model given by Eq. (5) have some guiding significance to describe the reaction mechanism. 0 The rate constant kf satisfies Arrhenius Eq. (7). 0
ln kf ¼ ln A
Ea RT
ð7Þ 0
Fig. 8 displays relationship between temperature and kf given in Table 5. We can see that the results are well fitted by Arrhenius equation. The activation energy is 17.74 kJ/mol and the frequency Table 5 Comparison of rate constants (present study: 0.85 MPa, Joelianingsih et al. [12]: 101.3 kPa). Present study
Fig. 6. Effects of methanol feed rate on the reaction conversion (290 °C, 0.85 MPa).
Joelianingsih et al. [12]
T (°C)
0 kf
230 250 270 290
0.0299 0.0375 0.0405 0.0481
1
(min
)
R2
T (°C)
kf (min1)
0.954 0.974 0.966 0.961
200 225 250
0.0052 0.0073 0.0095
0
379
H.J. Cho et al. / Fuel 93 (2012) 373–380 Table 6 Values of RMSE for each component (103). Temperature(°C)
FA FAME Overall
230
250
270
290
0.1792 0.3474 0.2682
0.2787 0.2058 0.2376
0.2637 0.2046 0.2290
0.3197 0.2281 0.2694
Fig. 8. Arrhenius plot for the esterification of PFAD (0.85 MPa, methanol feed rate: 2.4 g/min).
while the activation energy by Joelianingsih et al. [12] (24.8 kJ/ mol) is a little larger than present value. This means that the non-catalytic reaction is accelerated as temperature increases. 4. Conclusions A single-step non-catalytic esterification reaction using PFAD feedstock is investigated in this study. The PFAD feedstock contains 85–95% of various fatty acid components and resulting acid value satisfies the biodiesel product criterion of 0.5 (mg KOH/g). Reaction conditions are selected so that water being generated during the reaction is removed rapidly (temperature: >250 °C, pressure: 0.85–1.20 MPa, reaction time: <180 min). Effects of reaction temperature, pressure and the methanol flow rate are analyzed and reaction rate constants are determined by the 1st-order model of FA concentration. The concentration profiles of FA and FAME computed using aforementioned rate constants and reduced 1st-order model fit well with experimental data. From rate constants and Arrhenius relation, the activation energy and the frequency factor were found to be 17.74 kJ/mol and 2.12 min1, respectively. The main characteristic of the proposed reaction method is the employment of single-step reaction without catalyst to yield commercial grade biodiesel product with little loss. For this reason, the present reaction technique can be readily commercialized. Moreover, even low grade waste fatty materials such as waste greases can be used as feedstock, which was not permitted in traditional transesterification reaction due to high fatty acid content. References Fig. 7. Concentrations of FA and FAME: comparison between computational results (based on the 1st-order model) and experimental results (a) 230 °C, (b) 250 °C, (c) 270 °C, and (d) 290 °C.
factor is 2.12 min1. From comparison with values by Joelianingsih et al. [12], it appears that the frequency factors are about the same
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